Catalytic hydrocarbon conversion



ec., 1% E95@ J. A. RIDGWAY, JR., ETAL. 2,533,284

CATALYTIC HYDROCARBON CONVERSION Dec H2, E950 J. A. R1DGWAY,JR., ETAL 25339234 OATALYTIO HYDROOARBON OONvERsoN Filed Dec. 27, 194e 2 sheets-sheet 2 D l E 2o I0 20 30 40 50 60 70 80 90 IOO TOTAL CONVERSION W1". 7,

atentecl ec. l?,

UNITED STATES PATENT orties CATALYTIC iiYiRocAnBoN CONVERSION .'olm A. Ridgway, Jr., Texas Tex., and Philip- Hill, Hammond,- Inde, ass'gnors to Pari Al'ne'ri' can Renning- Corporation, Texas City, Tex.; a

corporation of Delaware application necesitar 27, 194.6, serial No. 718,817" 4 claims. (C1. 19e-52) formation of carbon or carbonaceous compounds' of such a character as to absorb or destroy the catalyst. Another object of the invention is to provide a process for cracking heavy liydrocarl bon oils with a minimum production of dry gas and a maximum amount of condensable hydro; carbon gases of isoparainic structure. Still another object of the invention is to provide a hy; drocarbon cracking process which selectively 'con'- verts a major part of the oil to gasoline and tar suitable for heavy fuel oil with a high heating Value While producing a minor amount of fixed gases such as hydrogen,- methane and ethane. come vapparent from the description which follows.

This application is a continuation-in-part oi our U. S. Patent No. 2,454l615, led November l2, 1943.

`The invention is illustrated by a drawing which shows diagrammatically in Figure 1 an appa; ratus for carrying out the process; Figure 2 is a graph of data illustrating an advantage of the process.

Heretoforthe catalytic cracking of heavy oils has been beset' by the great difficulty of catalyst deterioration, some processes requiring high cata lyst replacement and others requiring frequent catalyst regeneration at considerable expense. 1n the case of a catalyst such as aluminum chloL ride employed in the well-known McAfee process,

the aluminum chloride sufiered rapid containination resulting in extensive aluminum chloride losses owing to catalyst degeneration; The cause of degeneration is generally traceable to the interaction of the heavy hydrocarbons and their breakdown products, with the aluminum chloride; forming more or less stable addition products or complexes as a sludge of no catalytic value; Many attempts have been madeto regenerate the aluminumpchloride but Without much commer- C515lCCeSSa--,

In the case of high temperature catalytic crack- Other objects of the invention will be:

ing pr'ssseiplyng refractory porous solid catalysts such as the .active metal oxides, silica,

alumina, rrialgresia;V titania,zirconi a and mix- V tures thereof@ catalyst deeerftienis ,Ver rees and regeneration is required frequently, e. g'. after a few minutes toa 'Afewu hours. This regeneration isusually accomplished in the case of the refrac l tory metal oxide catalysts'by combustion with* air or other oxidizing gas. In the regeneration operation, considerable quantities o1 carbonaceous matter ,accumulated on the catalyst iby breakdown of the oil are removed in the oxidaion.

When substantially anhydrous liquid hydrofluoric acid is employed as a cracking catalyst-1 .in substantial amounts at moderately elevated temperatures, the conversion" can be carried oiit 1 with substantially no carbon formation but with the formation of considerable tar land heavy fueL, 0,11. CUS'. tillation of the conversion residue, the hydrol i. e. high boiling conversion products.

fluoric acid catalyst, having form'eddnorpe'rmanent compoundsl therewith, can be recovered sub-r; stantially completelyfor 11euse in Athe process.;`4

The heavy conversion products produced by the action of thi's'catalyst are valuable uelols afterAi removal of therefrom. Onecf their characteristics is thehigh heat of combustion coupledV with low specicgravityhmalring them very suit;

able for fuel purposes. We have now found that these heavy oils synthesized in theA process, herl in referred to as tan have a valuablecatalyst' promoter effect in the HF conversion reaction." This ajjlllii is dlCld l tllls' atlll lill process andI more specically to recycling the tar to the reaction zone.

The following examples illustrate the conver-V4 sion of gas oil with liquid HF catalyst. The op eration was continuous; the catalyst and oil being contacted in the istv two runs in a vertical, lpacked tower. In the Aother tworunsthe reactbr was equipped with a motor-driven sirrer. The pressure inthe reactorwas maintained at`900 p. s. i. g. to insure liquid phase conditions. The charging stock was a virgin gas oil with the io'l-A lowing characteristics:

The cracking conditions and results are shown in the following table:

Component yields, octane ratings and cracking conditions employed Tower Reactor Stirring Reactor Conditions Run Run Run Run No. 1 No. 2 No. 3 No. 4

AverageReactor Temperature..F 320 325 325 325 Charge Rate: lbs. ges oil per hour.. .597 1.88 1. 32 l. 30 Relative weight velocity (lb. oil/ hr./lb. HF) 0. 14 0. 46 0. 58 0.70 Contact Time, Minutes 11 11 8 25 Total Charge HF to Unit, Lbs... 15. 5 12.8 11.5 12.0

2 2gveragc catalyst ago at time of sampling (lbs. gas oil/1b. HF), 1.0-

Tower Reactor stirring Reactor Products Run Run Run Run No. 1 No. 2 No. 3 No. 4

Dry Gas (Cs and lighter) 2. 6 0. 17 0. 07 0. 56 Excess Isobutanc 11. 8 0.43 0. 72 3.16

Total 100.0 100.0 100.00 100.00

OCTANE RATINGS 0N GASOLINE FRACTIONS (400 E. P. 0. R. V. P.)

ASTCM Motor Method:

lear 68. 9 69. 7 65. 3 69. 0 1 cc. TEL 79. 0 80. 3 77.0 79. 7 ASTM Research Method: Clear. 68.4 69. 4 66.5 70.1

From these results it will be noted that there is a decrease in yield of gasoline with an increase in the relative weight velocity.

In batch operations, the following data show the results obtained from two runs on virgin gas oil carried out in a shaking bomb:

. Run Run Condmons No. l No. 2

Per cent Cat. by Weight 295 193 Per cent Cat. by Volume... 234 158 Per cent Cat. by Mols... 2300 2160 Time, hours 4 4 Temperature, C 155-165 155-162 Max. pressure, lbs/sq. in 910 840 Run Run Products, wt. percent oi Charge No. 1 No 2 'rar 45. 2 42.1 Yield Data (Based on amount reacting):

Totalgas (Ca, C4, Cs percent and lighter) 38. 6 31.0

Gasoline (Cn-200 C.) 11.7 15.9

Gas-I-gasolme, percent.. 50.3 46. 9

` Tar, percent 49.7 53.1

These results also show the importance of maintaining a substantial portion of HF catalyst in the conversion reaction. Thus, the total amount of conversion (total reacting) in run 1 with 295 per cent of HF was 91 compared with 79.3 in run 2, in which the conditions were the same except that the proportion of catalyst was less, i. e. 193%. In general, it is desirable to employ an amount of catalyst equal to at leastl half the Weight of the oil treated and a catalyst-oil ratio in the range of 1 :1 to 3:1 is desirable.

The tar produced in these runs corresponds to 13.3% and 17 .9% respectively, based on the weight of the tar-HF phase. In general better results are obtained with higher tar concentration, e. g. above 15% in the HF phase, readily obtainable by recycling.

Inspection of the above results also shows that the increased proportion of catalyst employed in run 1 favors the production of isobutane, a valuable product for use in alkylation and other reactions. Altho the recovery of tar in run 1 appears to be slightly larger than in run 2, when correction is made for the increased conversion in run 1 (91.0 V. 79.3) the tar yield will be seen to be less with increased proportion of catalyst to oil. (Note yield based on percent reacting 49.7 v. 53.1.) The temperature employed in liquid HF cracking catalyst is unusually low for catalytic conversion reactions. In our process the temperature is usually about to 230 C., preferably about to 175 C. In case a promoter is used with the HF, the temperature may be even lower, e. g. as low as 50 C.

rI'he reaction time may vary over a considerable range depending primarily on the temperature and the ratio of catalyst-to-oil treated. Thus, in an example in which 193 of HF was used to convert virgin gas oil, the total gas yield obtained from virgin gas oil at 135 to 145 C. was only 2.4% at a reaction time of two hours and 7% atve hours reaction time. However, in another example with the same catalyst concentration but with a conversion temperature of 155 to 162 C., an increase in reaction time of from 11/2 to 4 hours increased the gas yield from 3.4 to 24.6 per cent.

In general, it appears that a conversion temperature of C. or above is necessary for appreciable gas production within a period of four hours or less.

The apparent low yield of gasoline in the batch runs must be corrected by adding thereto the C5 and most of the C4 hydrocarbons when comparing with the usual conversion data on gasoline of 9 to 12 pounds R. V. P. On this basis the gasoline yield in batch run 2 would be abou 44% based on amount reacting.

In carrying out the process, we prefer to operate in a continuous manner as illustrated in Figure 1 of the drawing. Referring to the drawing, a suitable feed stock, for example a virgin gas oil of 35 A. P. I. gravity having a boiling range of about 345 to 650 F., is charged to the system by line I0 leading to mixer I I, preferably at reaction temperature or above. Any suitable mixing device may be used such as an orifice mixer consisting of a series of orifice plates thru which the feed stock is forced to flow at high velocity. Liquid HF is charged to the system by line I2 and mixed with the hydrocarbon feed in mixer I I. The mixture then flows by line I3 to reaction chamber I4. The volume of the reaction chamber is sufficient to provide the desired reaction time, for example ten minutes to four hours depending on the temperature, character of the feed stock, etc. The amount of HF employed in the reaction zone is sufficient to maintain a sep.

arate liquid catalyst phase and it is desirable to employ a considerable -excess above that required to saturate the hydrocarbons.

It is not necessary to mix the HF and feed stock before introducing into reactor I4 but these stocks may be injected directly into the reactor thru separate lines if desired, contact between them lbeing obtained entirely within the reaction chamber. In most case the liquid HF phase within the reaction chamber will be lighter than the hydrocarbon phase and will occupy an upper position in the reaction chamber. In this case it is desirable to introduce the hydrocarbon feed into the upper part of the reaction chamber and allow it to flow downwardly therethru in Contact with the liquid HF phase therein. Where the conditions are such that the HF catalyst phase is heavier than the hydrocarbon stocks in the reaction chamber I4, it is preferred to introduce the feed stock at a low point in the reaction chamber and withdraw it at a point near the top.

The temperature of the reaction chamber I4 is maintained above about 100 C. and generally within the range of 125 C. to 200 C., a suitable temperature being about 150 to 160 C. Higher temperatures may be employed for short reaction periods. Any suitable heating means may be employed for the purpose, for example coil I5 supplied with hot oil from steam or other heating fluid. We can also preheat the HF catalyst and hydrocarbon feed stock in separate heaters, not shown, previous to introducing it to reaction chamber I4. In some cases it is convenient to supply a portion of the feed as a vapor in suincient amount to control the reaction'temperature.

The hydrocarbon stock and HF in reaction chamber I4 may be agitated to obtain the necessary contact between the two liquid phases to eilect catalytic conversion, altho in some cases the degree of contact obtained by distributing the feed stock in small streams above a less dense liquid HF layer thru which the hydrocarbon passes in small streams or droplets will provide sunicient contact. Mechanical agitators may be installed directly in the reaction chamber or agitation may be provided both internally and externally of the reaction chamber as indicated in the drawing. According to this method, the reaction -mixture is withdrawn by line I6 from the bottom of chamber I4 and conducted by pump I'I thru line I8 back to the top of the reaction chamber. By controlling the circulation rate, any desired degree of interspersion of the liquid catalyst phase and the oil phase may be obtained.

-A rcontrolled stream of reaction products 4is withdrawn by valved line I9 leading thru cooler 20 to separator 2I from which some of the lighter gaseous reaction products can be withdrawn by line 22. Liquid reaction products are conducted by line 23 to catalyst settler 24. Settler 24 is preferably a horizontal, cylindrical, elongated chamber thru which the reaction products flow continuously without agitation, thereby allowing the liquid HF phase associated with tar toseparate at the bottom of settler 24. We have found that on cooling the mixture of hydrocarbon oils in liquid HF, an inversion of the layers commonly occurs because of a higher volumetric temperature coefficient of HF. Thus, Where the HF may form the upper layer at a temperature of 160 C. it will usually be the lower layer below the hydrocarbon at a temperature of 20 Ito 40C. Thev catalyst layer containing heavy hydroline 25 and pump 26 to reactor I4 byline 21, or to mixer II by line 28 for the treatment of fresh amounts of hydrocarbon feed stock. If desired, however, a part or all of the stream may be diverted to tar separator 29 wherein HF is distilled or stripped from the tar, the HF vapor being withdrawn by vapor line 30 leading to condenser 3l whence it is recycled by line 32 and pump 26 to reactor I4. The tar from which the HF has been completely or substantially removed is Withdrawn from the system by line 33. In this stripping step, light hydrocarbons can also be removed with the HF and recycled.

We have found that the recycle of the tar or HF-soluble fraction to the reactor is important in increasing catalyst activity in the cracking reaction catalyzed by HF and we prefer to recycle sufiicient tar, usually in HF solution, to maintain a concentration of tar `in the liquid HF phase -above about 15 per cent by weight based on the weight of the catalyst phase in the reactor. Note that in the batch operations hereinabove the catalyst phase contained 13.3 and 17.9 per cent of tar respectively.

The hydrocarbon reaction products are Withdrawn from settler 24 by line 34 to stripper r35. The pressure is reduced at valve 36, preferably to about 50 to 200 p. s. i. In stripper 35, most of the dissolved HF of which relatively little is present in the oil phase from settler 24 is eliminated as an azeotrope with hydrocarbon gases and conducted away by line 3l to a condenser or other means for HF recovery. Recovered HF vapor may, if desired, be recycled to the reactor I4. The liquid hydrocarbon products are withdrawn from tower 35 by line 38 leading to neutralizar 39 wherein traces of HF remaining in the hydrocarbon stream are removed by an alkaline neutralizing agent such as sodium carbonate, sodium hydroxide, lime, etc., either solid or in solution or by adsorption with a suitable adsorbent for HF such as fullers earth, silica gel, bauxite, or one of the acid-adsorbing nitrogen-base resins employed in Water treating.v In the case of certain stocks and other types of operation, where the amount of HF remaining in the products Withdrawn from stripper 35 is very small, the neutralizer may be omitted entirely. Y

From neutralizer 39 reaction products pass by line 39a. to fractionator 40. Heat is supplied to the fractionator by reboiler coil 4! and reflux by i cooling coil 42. Light products including butanes and lighter hydrocarbons are distilled off thru line 43 leading to fractionator 44.

The principal liquid hydrocarbon products are withdrawn from the bottom of fractionator 40 by line 45 and heated in heater 46 which may be a pipe still for example. From heater 46 the liquid products are conducted by line 4l to fractionator 48 provided with gas oil and gasoline side strippers 49 and 50 respectively. A heavy fraction suitable for fuel oil or asphalt manufacture is withdrawn atl the bottom by line 5I. Heat required for reboiling in iractionator 48 is supplied by coil 52 while reflux cooling is supplied by coil 53. The gas oil fraction withdrawn rby line 54, if desired, .may be conducted back to rcactor I4 or mixer I I as feedstock for the process. The gasoline withdrawn by line 55 is essentially va heavy blending naphtha.

Pentane and lighter hydrocarbons are .Withdrawn from fractionator 48 by .vapor lineileading to fractionator 4 4 towhich -yapors from fractionatorilllmay also be conducted. Propaneand carbons in solution can be conducted .directly v:by `lighter :products are discharged by Vapor line :53

7, while the pentane is removed as the bottom stock by line 58 and butane is withdrawn by line 59 from side stripper 6D. The pentane and butane streams produced in the process consist largely of isopentane and isobutane. The isopentane is chiey valuable for blending in aviation fuels and other high knock rating gasoline. Both isopentane and isooutane may be subjected to alkylation with suitable olens, for example ethylene, proplyene or butylene, to produce alkylate gasoline, neohexane, isooctane, trptane, etc., very desirable constituents of aviation fuels.

The following data illustrate the importance of recycling tar to the reaction zone. They were obtained from two runs in one of which the HF- tar phase was recycled. In the other run the operation was once thru with respect to catalyst. None of the hydrocarbons was recycled in either run.

Operating conditions Neccleyst Average Reactor Temperature, F 350 350 Reactor Pressure, p. s. i. f 900 000 API Gravity-Gas Oil Charge 36. 9 36. 0 Charging Rate: Volume of Oil per Hour per Unit Reactor Volume 0.56 84 Relative Weight Velocity, Lbs. per Hour per Lb. oi Catalyst 1.1 1. f. Catalyst-Oil Ratio in Reactor. 4.2 3.0 Contact Time, minutes 13. 4 12. 0

Yields on charge, Weight percent Products Catalyst No catalyst recycle recycle Dry Gas (C3 and lighter) 0.8 0.2 Excess lsobutane 2. 5 0.0 Gasoline (400 E. P., 10.() RVF) 33.0 15.8

Gas Oil:

v Hydrocarbon phase 40.5 62.5 HF phase 0.0 12.9

Total 4.6. 5 75. 4

'lar boiling above charge 17. 2 S. 6 Gas Oil Conversion 53. 5 24.6 gasoliscbutane)/tnr 2.06 1. 94

KNOCKING CHARACTERISTICS OF GASOLINE (400 E. P. 10.0 RVP) ASTM Motor Method:

It will be observed from these data that whereas the cracking conditions were substantially the same in the two runs, the contact time being only slightly longer in the first run, the yield of gasoline in the rst run was approximately double that obtained in the second run, i. e. 38.3% by volume in comparison with 19.1%.

These data show that the recycle of HF-soluble hydrocarbons from the reaction products separated as an insoluble HF-tar phase has the effect of greatly increasing the cracking rate. In addition the data suggest that there is an actual decrease in the amount of hydrocarbons which are degraded to tar (52.1 parts of tar per 100 parts of gasoline produced vs. 54.3), indicating that the recycle of tar to the reactor and the maintenance of a higher tar concentration therein serve to reduce tar formation. This effect is also shown in the curves in Figure 2 in which the overall tar yield is plotted against total conversion. The upper curve corresponds to the tar produced in a number of runs in which no tar was recycled but only the amount of tar produced in the con-.- version was present during the reaction. The lower curve corresponds to data obtained from a series of runs in which the charging stock, gas oil, was continuously converted in a reactor to which the HF tar phase was recycled from the product separator in accordance with the operation shown diagrammatically in Figure 1. 1t will be noted that the amount of tar produced from a given amount of charging stock is considerably less in the case where tar is recycled to the reactor increasing the concentration of tar in the HF catalyst phase.

As indicated hereinabove, one of the important advantages of this process of hydrocarbon conversion over other catalytic processes lies in the substantially complete recovery of catalyst without necessity of regeneration. Most of the catalyst separates as a separate liquid layer while the remainder is removed by simply stripping or distilling the HF from the reaction products. Substantially no undecomposable sludge or catalyst complex is formed in the reaction and therefore very little fresh catalyst must be added by way of replacement beyond that necessary to compensate for mechanical losses.

The conversion action of liquid HF may be modied if desired by adding small amounts of certain other reagents or promoters; for example by adding BFa the product distribution may be substantially altered. Amounts used are generally only about l to 10%. When using such promoters it is desirable to either recover the promoter along with the HF or if the promoter is discarded it is desirable to recover the HF therefrom, for example by distillation. In the case .of BFs it may be recovered from the gaseous hydrocarbon reaction products by means not shown in the drawing.

Hydrogen may also be employed in the HF conversion reaction to modify the character and amount of tar or asphalt formed. The amount employed can be in the range of 1000 to 3000 cubic feet per barrel of oil treated. Hydrogen pressures of 500 to 3000 p. s. i. are suitable. Hydrocarbon gases containing hydrogen may be used instead of hydrogen and hydrogen-containing gases produced in the process may thus be employed. In the HF catalytic cracking process, the gas produced consists chiei'iy of butanes and pentanes and in most cases these fractions contain about 88 to 93 per cent of isoparafdns. No neohexane has been detected. The butane yields are especially high; for example 22.4 per cent from virgin gas oil and 28.7 per cent from dodecene. The heavier charging stocks tend to produce less dry gas, that is propane and lighter hydrocarbons and hydrogen, than the lighter charging stocks. All gas is completely saturated. If desired, the C4 and C5 fractions may be allowed to remain with the gasoline instead of being separated as shown hereinabove. It is generally preferred to operate the process with sufficient cracking to yield products having a 90% point-v ASTM--below the 10% point of the charging stock.

The concentration of isobutane and isopentane is much higher than would be expected from the equilibrium values calcuiated from thermodynamic data, which may indicate that isomerization occurs prior to or simultaneously with the cracking reaction in the presence of HF. For comparison, the concentration of isobutane in the catalytic somerization of butanes is only about 65%.

The gasoline obtained in the process is usually substantially free of unsaturation. As shown in the table hereinabove, the octane number is about 70, and it is characterized by a high lead response, i. e. the addition of a small amount of tetraethyl lead increases the knock rating very materially as is characteristic with isoparaiinic hydrocarbons.

The tar produced in the process after the elimination of HF therefrom is characterized by a high specific gravity, for example about l to 1.15. A typical specimen of tar from a tar recycling operation had the following characteristics:

Density, D60/60 1.12 API gravity -5 Softening point (ball and ring) F 145 Penetration,1 200 gms. 60 secs. at 32 Fn 0 Penetration,1 50 gms. 5 secs. at 115 F 153 Penetration,1 100 gms. 5 secs. at 77 F-- 13 Ductility, 5 cms. per min. at 77 F 150+ Flash point (Cleveland open cup) F Solubility in carbon tetrachloride weight per cent 100 Solubility in n-heptane do 67.5 Solubility in acetone do 82.3 Iodine number (Wijs) I 140 Fluorine content weight per cent 0.05 Ultimate analysis:

Carbon weight per cent 91.9 Hydrogen do 7.9 Molecular weight (Menzies boiling point) 420-425 1 In 0.01 cm.

The excess tar not needed for recycling to the reaction zone can be employed as a heavy fuel oil or it may be converted into asphalt by heating and/or blowing with air, blending with other asphaltic stocks, etc.

Because of its high unsaturation it possesses drying properties making it useful as a component for coating compositions.

Altho we have described our process primarily as it is applied to the conversion of gas oil, it may also be applied to low knock rating heavy naphthas boiling, for example, in the range of about 325 to 450 F. where it is desired to obtain lighter hydrocarbons and particularly isoparafnic C4 and C5 hydrocarbons for alkylation and other special purposes.

In copending application for Letters Patent, Serial No. 731,744, led by Herschel D. Radford et al. on March 1, 1947, now Patent No. 2,527,573, there is claimed a process for cracking heavierthan-gasoline hydrocarbons with liquid hydrogen z a temperature above about 100 C. and under sufficient pressure to maintain a liquid catalyst phase in said reaction zone, gasoline is produced in said reaction zone concurrently with the production of unsaturated tarry products soluble in said catalyst phase, withdrawn from said reaction zone and separated from catalyst, the improvement comprising maintaining within said. reaction zone an amount of said tarry products substantially in excess of the amount produced in once-thru conversion of said hydrocarbon oils.

2. The process of converting a heavy hydrocarbon oil into gasoline and Similar lower boiling products which comprises introducing a stream of a heavy hydrocarbon oil boiling above gasoline into a reaction zone containing a body of liquid hydrofluoric acid catalyst, maintaining the temperature of the reaction zone in the range of about 100 to 300 C., maintaining suicient pressure in the reaction zone to provide a liquid catalyst phase therein, retaining said body of catalyst within said reaction zone and intimately contacting said oil with said catalyst, thereby substantially converting it into gasoline and a heavy tar higher boiling than said heavy hydrocarbon oil charged and soluble in said catalyst phase, separating gasoline from the catalyst and tar and withdrawing it from the system and maintaining the concentration of tar in said catalyst phase above about 15% by weight.

3. The process of claim 2 wherein the concentration of tar in the catalyst phase in said reaction zone is controlled by withdrawing a portion of the catalyst phase therefrom and subjecting it to stripping and reduced pressure, thereby removing HF from the tar component thereof and recycling said removed HF to said reaction zone.

4. The process of manufacturing gasoline by cracking a heavier hydrocarbon oil which comprises subjecting it at an elevated temperature to the action of a catalyst consisting essentially of a solution of tar and liquid HF, whereby said heavy hydrocarbon oil is substantially converted into gasoline and high-boiling tar, separating gasoline-containing hydrocarbon products from the tar-HF catalyst phase, separating tar from the catalyst phase and withdrawing it from the process, recycling HF to the conversion reaction and recycling sufficient tar to the conversion reaction to maintain a high speciic gravity of about 1 on said tar withdrawn from the process.

JOHN A. RIDGWAY, JR.

PHILIP HILL.

REFERENCES CITED The following references are of record in the le of this patent:

UNITED STATES PATENTS Number Name Date 2,357,495 Bloch Sept. 5, 1944 2,378,762 Frey June 19, 1945 2,405,993 Burk Aug. 20, 1946 2,416,184 Lee et al. Feb. 18, 1947 g 2,425,559 Passino et al Aug. 12, 1947 2,449,463 Evering et al Sept. 14, 1948 2,454,615 Ridgway et al Nov. 23, 1948 

1. IN THE PROCESS OF MANUFACTURING GASOLINE BY CRACKING HEAVIER HYDROCARBON OILS IN THE PRESENCE OF A CATALYST CONSISTING ESSENTIALLY OF HYDROGEN FLUORIDE, WHEREIN SAID OIL IS INTIMATELY CONTACTED WITH SAID HYDROGEN FLUORIDE IN A REACTION ZONE AT A TEMPERATURE ABOVE ABOUT 100*C. AND UNDER SUFFICIENT PRESSURE TO MAINTAIN A LIQUID CATALYST PHASE IN SAID REACTION ZONE, GASOLINE IS PRODUCED IN SAID REACTION ZONE CONCURRENTLY WITH THE PRODUCTION OF UNSATURATED TARRY PRODUCTS SOLUBLE IN SAID CATALYST PHASE, WITHDRAWN FROM SAID REACTION ZONE AND SEPARATED FROM CATALYST, THE IMPROVEMENT COMPRISING MAINTAINING WITHIN SAID REACTION ZONE AN AMOUNT OF SAID TARRY PRODUCTS SUBSTANTIALLY IN EXCESS OF THE AMOUNT PRODUCED IN ONCE-THRU CONVERSION OF SAID HYDROCARBON OILS. 